Fluid catalytic cracking processes and means

ABSTRACT

A PROCESS FOR CRACKING A PRINCIPAL HYDROCARBON CHARGE (SUCH AS A FURNACE OIL OR GAS OIL) CAPABLE OF BEING CRACKED TO GASOLINE IN THE PRESENCE OF A FLUIDIZED CRACKING CATALYST, SAID PROCESS COMPRISING THE STEPS OF MAINTAINING A PREDETERMINED RANGE OF TEMPERATURES WITHIN SAID CATALYST STREAM, ADDING A NAPHTHA DILUENT (GASOLINE TYPE HYDROCARBBON) TO SAID CATALYST STREAM, CONTROLLING THE PARTIAL PRESSURE OF SAID CHARGE IN SAID STREAM BY MAINTINING A GIVEN RATIO OF SAID DILUENT TO SAID CHARGE, AND ADDING SAID DILUENT TO SAID CATALYST STREAM AT A POINT HAVING A HIGHER TEMPERATURE THAN THAT AT WHICH SAID CHARGE IS ADDED SO THAT A SIGNIFICANT PROPORTION OF EACH OF SAID NAPHTHA AND SAID CHARGE IS CRACKED BY SAID CATALYST. AN EXEMPLARY CRACKING PLANT IS PROVIDED FOR EFFECTING THE PROCESS AS OUTLINED. METHODS ARE PROVIDED FOR THE UPGRADING OF VIRGIN AND CRACKED NAPHTHAS AND FOR ENHANCING THE OLEFINIC YIELDS THEREFROM.

Dec; 19, 1972 Filed Nov. 12,

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FLUID CATALYTIC CRACKING PROCESSES AND MEANS y Filed Nov. l2, 1969 .2Sheets-Sheet 2 `United States Patent ce Patented Dec. 19, 1972 3,706,654FLUID CATALYTIC CRACKING PROCESSES AND MEANS Millard C. Bryson, Conway,and Joel D. McKinney, Indiana Township, Beaver County, Pa., and .lamesR. Murphy, Huntington Station, N.Y., assignors to Gulf Research &Development Company, Pittsburgh, Pa.

Filed Nov. 12, 1969, Ser. No. 875,830 Int. Cl. C10g 11/14 U.S. Cl.208-74 25 Claims ABSTRACT F THE DISCLOSURE A process for cracking aprincipal hydrocarbon charge (such as a furnace oil or gas oil) capableof being cracked to gasoline in the presence of a uidized crackingcatalyst, said process comprising the steps of maintaining apredetermined range of temperatures within said catalyst stream, addinga naphtha diluent l(gasoline type hydrocarbon) to said catalyst stream,controlling the partial pressure of said charge in said stream bymaintaining a given ratio of said diluent to said charge, and addingsaid diluent to said catalyst stream at a point having a highertemperature than that at which said charge is added so that asignificant proportion of each of said naphtha and said charge iscracked by said catalyst. An exemplary cracking plant is provided foreffecting the process as outlined. Methods are provided for theupgrading of virgin and cracked naphthas and for enhancing the olefinicyields therefrom.

The present invention relates to the cracking of petroleum hydrocarbonfeed stocks to gasoline in the presence of a highly active fluidcracking catalyst and to the unexpected upgrading and cracking ofcertain components added with the feed stock.

In the aforementioned application, S.N. 836,383, a method is presentedfor improving the operation of a uidized, catalytic cracking processWithout lowering the pressure in the reaction zone or in the catalystdisengaging or stripping vessel. An unexpected advantage results fromcharging a diluent gas to the inlet of the cracking reaction zone tolower the partial pressure of the charge hydrocarbon in the system. Anydiluent which is a vapor or becomes a vapor under the conditions of thereaction zone can be used. -An inert gas such as steam or nitrogen is asuitable diluent. A mixture of gases can be employed. If the diluent isa hydrocarbon, it should desirably have a boiling point below about 430F., i.e., it should be a gasoline range hydrocarbon or lighter. If itboils above the gasoline range, it will, of course, itself be a portionof the cracking feed. Recycle methane or ethylene could be employed. Infurther accordance with the aforementioned application it was found thata lower hydrocarbon feed partial pressure at any given reaction zonetotal pressure produces the unexpected effect of increasing theselectivity to gasoline production at a given conversion level of freshfeed or, conversely, requiring a lower conversion of total feed toproduce a given gasoline yield.

In accordance with the aforementioned previous invention, it was furtherdiscovered that the gasoline selectivity advantage is transient and islost if the cracking process is not terminated in a timely manner, asexplained more fully in the first-mentioned previous application.Briefly, the amount of steam or other inert diluent must be sufficientto produce a significant reduction in partial pressure of the incominghydrocarbon feed. In general, the greater the amount of steam or otherdiluent gas introduced relative to the hydrocarbon, the greater will bethe effect upon selectivity, i.e., the greater the reduction in partialpressure, the greater the gasoline selectivity advantage.

In further accordance with the previous invention, it has beendiscovered that the gasoline selectivity advantage, owing to thepresence of an inert diluent which is not itself capable of beingcracked to gasoline, is most significant in the very early stages of thecracking reaction, which is also the period in which most of thecracking of fresh feed occurs. In point of fact, the productional curveof cracked hydrocarbon vapors from fresh feed with time is exponential,with the greatest rate of cracking occurring at the outset of thereaction so that the cracked vapors themselves quickly reduce thepartial pressure of the unreacted feed. However, by the time thesevapors are produced, most of the cracking has been completed. Forexample, after the hydrocarbon feed has been in the reaction zone forabout 0.1 second, it is about 40% converted and after 1.0 secondconversion has increased only to about to 80% In contrast to what wasknown before the firstmentioned Bryson and Murphy invention, reactionzone residence time is established not only by establishing the totalcharge rate, including both hydrocarbon and steam, but also byestablishing the ratio of steam or other diluent to hydrocarbon in thecharge. It was determined in accordance with the previous Bryson andMurphy invention that control of the ratio of diluent to hydrocarbonfeed or charge and control of the total charge rate including bothdiluent and hydrocarbon are interdependent and interdependently exert acritical effect upon a gasoline yield. The residence time of freshhydrocarbon feed in the reactor is controlled such that the maximumselectivity of the process for gasoline yield is not impaired by thephenomenon known as aftercracking Avoidance of aftercracking and also ofthe phenomenon known as backmixing are explained hereinafter withreference to FIGS. 1 and 3, and is further detailed in theaforementioned copending application.

In any particular cracking process the gasoline yield and residence timevalues which encompass the gasoline selectivity advantage of the presentinvention will depend upon many variables peculiar to a given process.Examples of these variables are the particular catalyst which isemployed, the level of carbon on the regenerated catalyst, catalystactivity and/ or selectivity, reaction temperature, and the refractorycharacteristics of the feed. The extent of the selectivity advantage ofour present invention might be as low as one-half percent to one percentor as high as three to five percent depending upon the desired ratio ofdiluent vapor to hydrocarbon feed at the reactor inlet. Considering thatgasoline is the most economically desirable product of the crackingoperation, the economic value of a selectivity advantage of evenone-half or one percent actually recovered as product effluent isconsiderable in a commercial reactor unit which can process 100,000 to150,000 barrels per day of hydrocarbon feed.

The present invention relates to a method for unexpectedly improving theselectivity to gasoline production in cracking processes utilizing auidized zeolitic cracking catalyst or a catalyst of comparable activityand/or selectivity. The present invention also provides methods andmeans for unexpectedly upgrading existing naphtha or gasoline stockthrough the cracking or analogous reaction. In this connection asurprising result is thereby evident from our invention, as the crackingreaction conventionally is used to produce gasoline stocks in the firstinstance, but not for the upgrading of existing gasolines.

The present invention, as in the case of the previous Bryson and Murphyinventions, is particularly useful in connection with cracking processesutilizing fiuidized catalysts. Natural or synthetic zeolitealuminosilicate cracking catalysts exhibit high activity in the crackingof hydrocarbon oils both in terms of total conversion of feed stock andin terms of selectivity towards gasoline production.

Although zeolitic aluminosilicates are especially useful catalysts forthe purposes of the present invention, any silica alumina or othercracking catalyst which is sufficiently active and/or selective to becapable of producing a transient maximum or peak gasoline yield from thetotal fresh hydrocarbon feed capable of being cracked to gasoline atresidence times of seconds or less are within the purview of our presentinvention. The maximum gasoline yield obtained at residence times within5 seconds is transient and rapidly diminishes. After a residence time ofone second, most of the fresh hydrocarbon feed is converted, and thereis a sharp drop in rate of conversion of fresh feed. However, if thehydrocarbon continues to remain in contact with the catalyst, productsof the earlier cracking operation themselves in turn undergo cracking.This occurrence is termed after-cracking. Since there is a greaterabundance of cracked material than uncracked material after only aboutone-half to one second of reaction zone residence time the situationrapidly arises wherein considerably more cracking of cracked thanuncracked material can occur. When this situation prevails, the desiredgasoline product initially produced at a high selectivity in accordancewith the present invention becomes depleted owing to aftercracking at afaster rate than it is-replenished from cracking of remaining uncrackedfeed so that the selectivity advantage initially achieved issubsequently lost at a significant rate. If timely disengagement ofhydrocarbon and catalyst does not occur prior to the occurrence of asignificant amount of aftercracking the very existence of the earlieradvantageous selectivity effect can be entirely masked.

In fiuid catalytic cracking operations it is generally advantageous tooperate the cracking reactor at pressures in the range of about to 30pounds per square inch gauge and it is undesirable in terms of theintegrated operation, including catalyst regeneration and power recoveryfrom regenerator flue gases, for reactor pressures to fall significantlybelow this level. For example, catalyst regeneration is generallyfavorably influenced by elevated temperatures and pressures.Furthermore, in systems where regenerator flue gas is utilized to dri-vea turbine to compress combustion air to be supplied to the regenerator,it is important to maintain an elevated pressure in the regenerator inorder to obtain efiicient turbine operation. Since spent catalyst mustfiow from the reactor zone to the regenerator, a correspondingly highpressure is consequently required in the reactor in order to urgecatalyst towards the regenerator. However, as noted previously anddescribed more fully below, relatively high reactor hydrocarbon feedpressures are less favorable to gasoline selectivity in the crackingoperation than are relatively low pressures.

It has been known, even before the advent of the previous Bryson andMurphy invention, that the use of an inert diluent such as steam at thehydrocarbon feed zone accomplishes certain advantageous effects in afiuidized catalytic cracking operation, such as assisting infiuidization of the catalyst, vaporization of liquid feed, dispersal ofcatalyst into the hydrocarbon feed, and increasing the reaction rate.However, such prior art provides no teaching of the proper control ofdiluent quantity and feed rate vis-a-vis the hydrocarbon feed in orderto obtain the aforementioned improvement in gasoline selectivity. lt waspreviously considered that the amount of steam to be employed in a fluidcatalytic cracking process should not be great in order to avoid areduction in residence time and thereby a loss in conversion rate. Thisis, of course, contrary to the findings of the previous Bryson andMurphy invention and also of our present invention. In the controlmethods for fiuidized catalytic cracking operations, as they existedprior to our aforementioned co-invention, a vapor such as steam wasadded to the inlet of the reaction zone, Which may comprise an elongatedriser tube, to assist dispersal of the catalyst into the hydrocarbonfeed. The amount of steam was not considered particularly critical.Reaction residence time (space velocity) was then adjusted to controlgasoline yield in the reactor efiiuent. If analysis of reactor effluentindicated that an adjustment of the residence time Was required, thehydrocarbon fiow rate was adjusted. However, no criticality was attachedto the fact that this adjustment also varied the ratio of steam tohydrocarbon at the reaction zone inlet. In short, the partial pressureof the hydrocarbon feed within the reaction zone was left largely tochance.

The invention disclosed in the first-mentioned copending applicationoffers, then, a significant improvement over the prior art, in theserespects, as briefly described above. The previous Bryson and Murphyinvention teaches the proper control of hydrocarbon partial pressure andresidence time within the reactional zone in order to maximize gasolineyield and selectivity. Control of partial pressure and residence timesare further elaborated upon below with reference to FIGS. 1 and 2, whichprovide also a basis for our present invention. The improvement ingasoline yield as a result of the previous Bryson and Murphy inventionis significant and should not be minimized. However, this previousinvention contemplates the addition of a diluent material which is a gasor vapor at the reactor conditions attendant upon its entrance, butwhich is largely inert under such conditions. That is to say, thediluent does not itself significantly enter into the reaction butprovides only the primary function of reducing the partial pressure ofthe hydrocarbon feed. This remains the case whether steam, nitrogen,methane, ethylene, naphtha, or other hydrocarbon having a boiling pointbelow about 430 F. is utilized as the diluent and is added concurrentlywith the hydrocarbon feed. Even if naphtha or similar lower boilinghydrocarbon were used pursuant to the teachings of the previous Brysonand Murphy inventions, the diluent thus comprised is essentially inertat the temperature and reaction conditions attaining in the reactionalzone downstream from its point of entry. Under these conditions, naphtha(boiling between and 430 F.) is suiciently refractory to avoid anyconversion of the naphtha. Although the previous inventions contemplatea further lowering of the hydrocarbon feed partial pressure while in thereaction zone by virtue of cracking of the gas oil, a significantlowering of the partial pressure from this source is delayed until afterthat residence time corresponding to the maximum selectivity range.

The present invention contemplates an unexpected further lowering of thevapor pressure of the gas oil or similar hydrocarbon feed by providingunique methods and means for the prior cracking of a naphtha diluentbefore admixture with the principal hydrocarbon feed.

Present trends in the petroleum industry have caused concern overpossible shortages of light olefins. Although the aforementionedzeolitic aluminosilicates have demonstrated substantial increases ingasoline conversion, they tend to yield less light olefins and otherlight ends than the previously used amorphous silica-alumina catalysts.This applies with equal force to the aforedescribed, as

Well as other prior catalytic cracking processes which employ zeoliticcracking catalysts. Concurrently with the present day downward trend inthe production of light oleiins, light olens are experiencing anincreased demand throughout the petroleum industry. The light oleiinicends find a most important use as sources of high octane alkylate forgasoline blending stocks, as well as for general petrochemical usage.

Our invention, therefore, further contemplates cracking of a naphthastock in conjunction with the usual hydrocarbon feed in a catalyticcracking unit to obtain, unexpectedly, a higher yield of light olefinicends. In this connection we have further discovered that the unconvertedportion of the naphtha feed is unexpectedly upgraded from lower octanegasoline constituents to constituents having a significantly higheroctane rating.

It is known, of course, that either a virgin or pyrolytic naphthafraction can be cracked to obtain light olens and for other purposes.The process, however, heretofore has been generally nnrewarding as thenaphtha fraction itself includes gasoline constituents, and, previously,there has existed no motivation for the charging of naphtha as feedstock into a catalytic cracking unit. In the usual or known crackingoperation gasoline constituents are, of course, produced by the crackingof larger molecules. No process is known to us for the upgrading ofexisting gasoline constituents either alone or in conjunction with theusual hydrocarbon feed stocks to improve the octane rating of thenaphthas. In point of fact, feeding of gasoline constituents into aconventional catalytic cracking unit has been carefully avoided in thepast whenever possible. Further, the cracking of naphtha alone not onlydoes not result in a significant upgrading of the gasoline rating, butis dicult to sustain on a continuous basis as the cracked naphthadeposits very little coke or residual carbon upon the uidized catalyst.Although our present invention takes advantage of this fact, naphthacracking by any of the known processes leaves an insufficient deposit ofcarbon on the catalyst to supply the heat required for reheating andregenerating the catalyst in the regenerating vessel.

yWe overcome these disadvantages of the prior art and accomplish theaforementioned desirable results by providing a fluidized catalyticcracking process wherein a hydrocarbon feed and diluent naphtha areadded separately to adjacent reaction zones or at spaced locationswithin the same cracking reaction zone. The naphtha diluent, which ismore refractory than the hydrocarbon feed at reaction temperatures ofabout l000 F. and lower, preferably is added at a location of highertemperature so that it undergoes a significant cracking reaction priorto mixture with the hydrocarbon feed. Addition of the diluent naphtha iscontrolled such that the attendant ternperature drop of the fluidizedcatalyst does not fall below that required for eicient vaporization andcracking of the gas oil charge. The respective points of introduction ofthe diluent naphtha and the hydrocarbon are sufficiently separated sothat significant cracking and upgrading of the diluent naphtha occursduring the residence time of the flowing catalyst and diluent streamsbefore the main hydrocarbon feed stream is encountered with attendantvaporization and cracking of the naphtha diluent leaves very littleresidual carbon upon the uidized catalyst, and thus its ecacy forsubsequent cracking of the hydrocarbon feed is substantially unimpaired.The vaporization and cracking of the diluent, moreover, affordsadditional mols of gas per unit weight of added diluent to provide afurther and more advantageous lowering of the main hydrocarbon feedvapor pressure, with a given quantity of the diluent.

With our novel method, as briefly described above, we have been ableunexpectedly to upgrade the quality of the unconverted naphtha derivedfrom constituents of either a virgin or pyrolytic naphtha fraction, toincrease the olefinic yield thereof, to lower advantageously the partialpressure of a hydrocarbon feed, to control the residence times of boththe naphtha diluent and the hydrocarbon feed, and to increase thegasoline selectivity based upon a given hydrocarbon charge.

More specifically, in accordance with the present invention a preheatedliquid hydrocarbon charge and a preheated uidized zeolite or comparablecracking catalyst are added to a primary cracking reaction zone. Thepartial pressure of the hydrocarbon charge is established principally bythe previous addition of diluent naphtha to the iluidized catalyst asdescribed below. Further adjustment of the charge hydrocarbon partialpressure can be accomplished, if desired, by the concurrent addition ofan inert diluent either directly to the primary reaction zone, to thecharge inlet line, or to both. The character of the latter diluent,which can be steam, nitrogen, recycle meth-. ane or ethylene, etc., issuch that, when used, it is gaseous under the conditions prevailing inthe primary reaction zone.

In further accordance with the invention, a preheated diluent naphthacharge (boiling Ibetween about 100 F. and 430 F.) is added to thefluidized catalyst at a secondary cracking reaction zone, which,however, is disposed upstream of the primary reaction zone. At the pointof its introduction into the secondary reaction zone, the iluidizedcatalyst stream is, therefore, at a substantially higher temperature.The naphtha diluent is substantially instantaneously vaporized and asubstantial proportion thereof is almost immediately cracked and/ orupgraded to low-boiling gasoline constituents, or cracked to the lighterolens and other light ends. At the downstream end of the secondaryreaction zone, the temperature of the combined tiuidized catalyst anddiluent stream has dropped to that which is appropriate for selectivelycracking the principal hydrocarbon charge (usually a gas oil boilingbetween about 430 and ll00 F.) the principal hydrocarbon charge islikewise instantaneously vaporized upon its introduction into theaforementioned, primary reaction zone. The quantity of cracked naphthadiluent at the entrance to the primary reaction zone is sufficient toaccomplish a substantial reduction of the partial pressure of thehydrocarbon charge. Cracking of the diluent in the preliminary orsecondary reaction zone further lowers the partial pressure of theprincipal charge. As much as about 80% of the naphtha diluent can becracked in the preliminary zone, which enhances the diluent function ofthe naphtha. Any required adjustment during the continuous crackingoperation can be effected by varying the amount of naphtha diluent addedto the upstream, secondary reaction zone, or alternatively, or inconjunction therewith, by adding or varying an inert gaseous diluent,which can 'be supplied at the inlet of or upstream from the downstreamor primary reaction zone as noted above.

Owing to the reactive conditions to which the diluent naphtha issubjected by our novel process, a very wide range of naphthacompositions or gasoline fractions can be employed. Because our novelprocess exhibits increased selectivity to oleiin production,particularly among the lighter ends, a naphtha having a preponderance ofparaflins and other hydrocarbon saturates can be used. For this reason,a naphtha constituted of at least 70% by volume of saturates, andpreferably higher, is a desirable diluent, as such compositions are lessdesirable for other purposes and yield an increased olenic production.For the same reasons, a naphtha composed of 10% by volume or less, andpreferably 5% or less, of oleiins is desirable.

The selectivity to gasoline production of the cracking process isenhanced owing to the lower partial pressure of the principalhydrocarbon feed at the entrance of the primary reaction zone, wroughtby the cracked naphtha diluent from the upstream or secondary reactionzone either alone or in combination with an inert diluent added at theinlet of or upstream from the downstream or primary reaction zone. Toprevent subsequent loss of the selectivity advantage, `both the diluentcharge and the principal hydrocarbon charge are permitted to remain inthe presence of the catalyst only as long as further conversion ofuncracked hydrocarbon produces a significant increase in gasoline yield.The system is controlled so that substantially at the time when furtherconversion of uncracked hydrocarbon produces no significant increase ingasoline yield or at the time when some decrease in gasoline yieldensues, the catalyst and hydrocarbon are substantially instantaneouslydisengaged from each other to prevent aftercracking of gasoline product(derived primarily from the hydrocarbon feed but also and to a limitedextent from the uncracked portion of the naphtha diluent) from negatingthe selectivity advantage initially achieved owing to the partialpressure effect.

Analysis of the product to measure total conversion of both feed anddiluent, or gasoline yield or both will aid in controlling the crackingreaction in accordance with this invention. These analyses Will providea measure of gasoline selectivity for controlling the reaction. Reactiontime duration can be adjusted by regulation of `cyclic catalyst rate,principal charge rate, diluent input rate, and the ratios therebetween,where the respective lengths of the primary and secondary reactor zonesare fixed.

In accordance with this invention, the reactor is operated so that thereis a continual increase in gasoline yield throughout substantially theentire length of the reactor coupled with a corresponding decrease inthe unreacted proportion of the hydrocarbon feed. This permits thereaction to be terminated at or near the time of maximum gasoline yield.Significant Ibackmixing in the primary and secondary reaction zones isavoided, as this would lead to aftercracking. Backmixing can result froman excessive linear velocity and attendant turbulence, or by theformation of a dense catalyst bed which induces turbulence in theflowing vapors. The principal hydrocarbon charge and the diluent remainin the primary reaction zone only until a decrease in the proportion ofunreacted feed is not accompanied by any substantial net increase ingasoline yield. Maximum gasoline yield is accompanied by maximumgasoline selectivity.

The overall time of contact between principal hydrocarbon charge andcatalyst can be as low as about 0.5 second or less but not greater thanabout seconds and will depend upon many variables in a particularprocess such as the boiling range of the charge, the particularcatalyst, the amount of carbon on the regenerated catalyst, the catalystactivity, the reaction zone temperature, and quantity of polynucleararomatics. The reaction should be permitted to proceed long enough tocrack any monoor di-aromatics or naphthenes because their reactionproducts result in relatively high gasoline yields and are the mostreadily crackable aromatics, but the reaction should be terminatedbefore significant cracking of other polynuclear aromatics occurs, ascracking of these latter compounds occurs at a slower rate and resultsin excessive deposition of carbon on the catalyst. It is clear that nofixed cracking time duration can be set forth but the time will have tobe chosen within the aforesaid range depending upon the particularsystem. In one system, even slightly exceeding a 1.0 second residencetime might result in such severe aftercracking that the selectivityadvantage would be lost, while in another system unless a 1.0 secondresidence time is appreciably exceeded there might not be sufficientcracking of charge hydrocarbon to render the process economicallyadvantageous. Generally, the residence time will not exceed 2.5 or 3seconds and 4 second residence times will be rare.

We accomplish these desirable results by providing a process forcracking a principal hydrocarbon charge capable of being cracked togasoline in the presence of a fluidized cracking catalyst, said processcomprising the steps of maintaining a predetermined range oftemperatures within said catalyst stream, adding a naphtha diluent tosaid catalyst stream, controlling the partial pressure of said charge insaid stream by maintaining a given ratio of said diluent to said charge,and adding said diluent to said catalyst stream at a point having ahigher temperature than that at which said charge is added so that asignificant proportion of each of said naphtha and said charge iscracked by said catalyst.

We also desirably provide a similar process including the additionalsteps of establishing a ratio of said naphtha diluent and the conversionproducts thereof to said principal charge and of establishing aresidence time of said principal charge such that a greater percentageyield of gasoline based on total hydrocarbon feed is recovered from saidprocess in the presence of said naphtha diluent and its conversionproducts than could be recovered from said process in the absence ofsaid naphtha diluent.

We also desirably provide a similar process including the additionalstep of establishing the ratio of said naphtha diluent and itsconversion products to said principal charge such that a greaterpercentage yield of lighter olefins based on total hydrocarbon feed isrecovered from said process in the presence of said naphtha diluent thancould be recovered from said process in the absence of said diluent.

We also desirably provide a similar process including the additionalstep of establishing the ratio of said naphtha diluent to said principalcharge and a predetermined temperature and residence time of saiddiluent in said catalyst stream prior to engagement with said principalcharge such that the octane rating (quality) of the unconvertedproportion of said naphtha diluent is upgraded.

We also desirably provide a similar process including the modified stepof adding said naphtha diluent in the range of about 5 percent to about45 percent by volume of the total hydrocarbon feed.

We also desirably provide in a catalytic cracking plant, the combinationcomprising conduit means for defining a primary reaction zone and foradding a principal hydrocarbon feed thereto, additional conduit meansfor defining a secondary reactio-n zone upstream of said primary zoneand for adding a diluent hydrocarbon feed thereto, and catalyst conduitmeans connecting said zones for circulating a catalyst streamsuccessively through said secondary and said primary zones.

During the foregoing discussion, various objects, features andadvantages of the invention have been set forth. These and otherobjects, features and advantages of the invention together withstructural details thereof will be elaborated upon during theforthcoming detailed description of certain presently preferredembodiments of the invention and presently preferred methods ofpracticing the same.

In the accompanying drawings we have shown certain presently preferredembodiments of the invention and have illustrated certain presentlypreferred methods of practicing the same, wherein:

FIG. 1 is a graphical representation of the variation in unreactedcharge and gasoline yield versus reactor residence time, pursuant to oneaspect of my invention;

FIG. 2 is another graphical representation illustrating the effect ofvariation in partial pressure of hydrocarbon feed upon debutanizedgasoline yield and total conversion rate; and

FIG. 3 is a schematic apparatus and fluid fiow diagram of an exemplarycatalytic cracking operation arranged according to my invention.

A reference to FIGS. 1 and 2 will illustrate the significant improvementwrought by the present invention.

FIG. 1 contains curves semi-quantitatively relating the amounts ofunreacted charge and gasoline, as a percentage of fresh feed, toreaction zone residence time. Curve a of a unreacted principal charge,typical of most uid cracking charge stocks, shows that the amount ofunreacted charge (curve a) asymptotically approaches a value somewhatless than 2O percent of fresh feed within the residence timescontemplated by our novel process. The

gasoline curves show that the quantity of gasoline produced rapidlyreaches a somewhat rounded maximum or peak which generally coincideswith the time at which the cracking rate of unreacted charge becomessubstantially diminished. The gasoline yield at the peak for a givencharge will be determined primarily by reactor temperature, to an extentby the level of carbon on the catalyst, and to an extent by thecatalyst-to-oil ratio. After reaching a peak the gasoline leveldiminishes because the aftercracking of gasoline predominates overproduction of gasoline from the unreacted feed. The lower gasoline curveb shown in FIG. l indicates that level of gasoline which would attain inthe reaction zone, assuming substantially no diluent naphtha isintroduced. The upper gasoline curve c shows the higher gasoline levelachieved by adding the naphtha diluent to the cracking process to lowerthe principal hydrocarbon charge partial pressure and thereby toincrease selectivity to gasoline. A still larger gasoline yield,attainable with our present invention results when the increase in lightolefin yield, also provided by this invention, is converted intogasoline by alkylation.

To illustrate the advantages of controlling residence time and principalcharge partial pressure, it will be assumed that a fluid crackingprocess is opertaing with addition of diluent naphtha to the secondaryreaction zone (with or without direct diluent addition to the primaryreaction zone) and the gasoline yield is at point A shown in FIG. lWhere significant aftercracking has occurred. In order to reduce theextent of aftercracking, it is decided to increase the charge rate ofhydrocarbon into the primary reaction zone, thereby reducing thehydrocarbon residence time. In conventional processes, residence time isusually adjusted by changing the hydrocarbon charge rate rather thandiluent charge rate since for any given percentage increase or decreasein charge rate of diluent or hydrocarbon, the effect upon reactionresidence time will usually be much greater in the case of thehydrocarbon adjustment because the total amount of hydrocarbon chargedis usually much greater than the total amount of diluent charged. Owingto the shorter residence time and concomitant reduction inaftercracking, a higher gasoline yield B is achieved. However, becausethe hydrocarbon partial pressure at the primary reaction zone inlet hasbeen increased by an increase in hydrocarbon tiow rate, the point B isremoved from the upper gasoline curve c in the direction of the lowergasoline curve b and is outside the cross-hatched zone e (FIG. l) whichdenotes a range of improvement proffered by this aspect of theinvention. The cross-hatched Zone e denotes the transient elevatedgasoline yields which can be recovered by the use of diluent vapor orcombination of vapors but which could not be recovered in the absence ofsuch vapor or vapors. On the other hand, if the same decrease inhydrocarbon residence time were achieved by increasing both theprincipal hydrocarbon and diluent flow rates in the same ratio so thatthe partial pressure of the principal hydrocarbon charge at the reactionzone inlet remained unchanged at the new residence time, the newoperating point would be at B', instead of B, which is within theaforementioned range of diluent improvement. On the other hand, if thesame total flow rate were achieved by increasing the ratio of thenaphtha and naphtha conversion products (by prior cracking of naphthadiluent) to principal hydrocarbon charge the new operating point wouldbe above b', and of course the area covered by the cross-hatched zone Eof FIG. l would be enlarged. Now, if the hydrocarbon charge rate isagain increased in a conventional cracking operation, to further reduceresidence time, a point C is reached which is still further removed fromthe middle gasoline curve C in the direction of the lowest gasolinecurve B than is point B because the hydrocarbon partial pressure hasbeen further and disadvantageously increased in going from point B topoint C. Again, because of the increase in hydrocarbon partial pressure,point C is outside the optimum selectivity range of my novel process. Onthe other hand, if the same residence time indicated at point C isachieved by increasing the flow rates of both diluent and hydrocarbon,rather than hydrocarbon alone, so that the hydrocarbon partial pressureat the new residence time is the same as it was at point A, the point Cis achieved which is within the optimum selectivity range. Again, theprior cracking of naphtha diluent according to the present invention,results in a further lowering of partial pressure and attendently highergasoline selectivity.

It is seen from FIG. l, that the operating points B and C, provided byconventional processes, represent essentially similar gasolineconversion levels occurring at different residence times, and one mightreadily make the erroneous assumption that these points dene a flatmaximum gasoline yield, however, points B and C lie outside the range ofthe present invention whereas operating points B and C fall within theoptimized selectivity range of this invention. Thus points B' and C lieat higher gasoline yield levels than points B and C, even though pointsB and B and points C and C represent the same residence times,respectively. Starting from point A, point B is reached by the method oflowering residence time via a change in both diluent flow rate andprincipal hydrocarbon tlow rate while, also starting from point A, pointB is reached by the method of changing the hydrocarbon flow rate only toachieve the same residence time as point B. Starting from point B',point C is reached by changing both diluent flow rate and hydrocarbon owrange to lower the residence time, while point C is reached by thesimpler and conventional method of changing the hydrocarbon flow rateonly to achieve the same residence time as at point C. It is apparentthat to achieve the aforementioned gasoline selectivity advantage theresidence time and the apportioning of diluent and hydrocarbon flowrates to achieve said residence time are interdependent and represent acritical combination for purposes of process control.

In making the aforementioned changes in diluent ow rate to maintain aconstant partial pressure of the hydrocarbon charge or to reduce thepartial pressure thereof, it is within the scope of the presentinvention to vary the liow rate of the diluent naphtha which supplied tothe inlet of the secondary reaction zone for subsequent mixing of bothunreacted naphtha and cracked naphtha product with the hydrocarboncharge at the entrance to the primary reaction zone, i.e., where theeffluent from the secondary zone enters the primary reaction zone. It isalso contemplated that the total diluent added to the process cancomprise the aforementioned naphtha diluent plus a minor quantity of aninert diluent or additional naphtha diluent added directly to theentrance of the primary reaction zone or alternatively or in combinationtherewith added directly to the hydrocarbon feed stream prior to itsdelivery to the primary reaction zone. Thus, in accordance wtih myinvention, the diluent to hydrocarbon ratio can be modified by adjustingthe feed rate of naphtha diluent to the secondary reaction zone, byadjusting diluents (if used) supplied directly to the primary reactionzone, by adjusting diluents (if used) supplied directly to thehydrocarbon feed stream, or by a combination of two or more of these.

The reaction temperature in the primary reaction zone, in accordancewith this invention can range between about 900 F. and about 1l00 F.Desirably, the temperature range is maintained between 950 F. and 1000F. The total pressure in the primary reaction zone can vary widely andcan be for example 5-50 p.s.i.g. or preferably 20-30 p.s.i.g. The totalpressure in the preliminary or secondary reaction zone (where thenaphtha diluent is cracked in the absence of principal hydrocarboncharge) desirably is maintained within the range of 5 to 50 p.s.i.g. Themaximum residence time in the primary reaction zone is 5 seconds and formost charge-stocks, the residence time will be about 1.5-2.5 seconds inmost 11 cases or, less commonly, 3-4 seconds. For high molecular weightcharge stocks which are rich in aromatics, a 0.5- 1.5 second residencetime is suitable in most cases in order to crack monoand di-aromaticsand naphthenes which are the aromatics which crack most easily and whichproduce the highest gasoline yield, but to terminate the operationbefore appreciable cracking of polyaromatics occurs because thesematerials produce high yields of coke, C2 and lighter gases.

In order to minimize the deposition of coke on the catalyst in thepreliminary or secondary reaction zone, the maximum residence time ofthe combined catalyst and diluent stream therein, is limited to a rangeof about 2 seconds to about 20 seconds and preferably to 2-10 seconds.Limitation of the residence time in the secondary reaction zone in thismanner maximizes the conversion of the naphtha diluent but minimizes thereduction in catalytic effect of the zeolitic material when subsequentlyengaged to the principal hydrocarbon charge at the entrance to theprimary reaction zone. Limiting the residence time in the secondaryreaction zone also avoids aftercracking of the naphtha diluent andattendant production of C2 and lighter gases, and coke. The quantity ofnaphtha diluent, added to the entrance of the secondary reaction zonecan vary between about 5 and about 45 percent by volume (with about5-20% being preferred) based on the total hydrocarbon (gas oil plusnaphtha) charge. The ratio of diluent naphtha to primary hydrocarboncharge can be varied depending upon the desired extent of partialpressure depression of the primary charge. Preferably, the quantity ofdiluent naphtha added to the secondary reaction zone is limited to about45 percent by volume in order to provide an adequate residence time inthe secondary reaction zone and to limit the yields of coke and C2 andlighter gases. In accordance with the invention additional diluents,where desirable, can be added by injecting either a naphtha diluent oran inert such as one of those mentioned previously at the entrance ofthe primary reaction zone or preliminarily into the hydrocarbon feedstream.

The length to diameter ratio of the primary reaction zone can varywidely, but the reactor should be elongated to provide a high linearvelocity, such as 25-75 feet per second, and to this end a length todiameter ratio above or 25 is suitable. The primary reaction zone canhave a uniform diameter or can be provided with a continuous taper or astep-wise increase in diameter along the reaction path to maintain anearly constant velocity along the ow path. The amount of diluentsupplied to the secondary reaction zone or concurrently to both reactionzones can vary depending upon the ratio of primary hydrocarbon todiluent desired for control purposes.

In accordance with the invention, the temperature of the initiallyadmixed catalytic and naphtha diluent stream at the entrance to thesecondary reaction zone is maintained in the neighborhood of about 250F. higher than the fluid stream temperature adjacent the entrance of theprimary reaction zone. A desirable temperature range at this point inthe secondary reaction zone is 1200- 1250 P. although the secondaryreaction zone can be operated with a substantial degree of success inthe temperature range of about 1100-1300 F. The higher temperature rangein the secondary reaction zone is desirable to promote cracking ofeither virgin or pyrolytic diluent naphtha.

Cracking of the more refractory naphtha, which is occasioned by thehigher temperature range in the secondary zone, results in increasedmols of gases from a unit weight of naphtha diluent which further lowersthe partial pressure of the subsequently added, principal hydrocarbonfeed in the primary reaction zone, while minimizing the quantity ofnaphtha diluent initially added to the secondary reaction zone. Thehigher temperature maintained in the initial stages of the secondaryreaction zone also increase the selectivity of the cracking reaction inthe secondary zone to light olefins. As noted previously, an increasedolefinie production is extremely advantageous in maximizing the ultimategasoline production from cracking and attendant alkylation operations.Notwithstanding the higher cracking temperature in the secondaryreaction zone, the deposition of coke from the naphtha diluent upon thecatalyst is extremely low so that the eicacy of the catalyst forsubsequently cracking the principal feed is substantially unimpaired. Inthe primary reaction zone, at least half of the heat of the catalyst isimmediately taken up by vaporization of the hydrocarbon feed and theremainder is applied to cracking the feed. The heat of vaporization ofgas oil for example is about the same as the heat of the crackingreaction or about B.t.u. per pound for cracking in comparison with aboutB.t.u. per pound for vaporization. Thus, little heat is available forcracking or aftercracking of the naphtha diluent, when the latter entersthe primary reaction zone. The desirability of adding the naphtha andgas oil at spaced locations in thereactional system is thereby apparent.

The boiling point of the naphtha diluent can vary characteristicallybetween about F. and about 430 F. It follows that the naphtha diluentitself is in the gasoline boiling range. As a practical matter thatportion of the naphtha fraction having a lower octane rating or which isotherwise unsuitable for various reasons as a gasoline constituent ispreferably used as a diluent. For example, a lighter naphtha fraction (cg., one boiling between 100 F. and 290 F.) can be employed where theultimate gasoline blend is destined for warmer climates or seasons, inwhich blend the lighter naphtha fraction is not ordinarily desirable.Conversely, in the production of gasolines which are blended for use incolder climates or seasons, the heavier naphtha fraction (e.g., boilingbetween 290 F. and 430 F.) is more desirable for diluent purposes, asthe aforementioned lighter fraction then becomes more suitable as agasoline constituent.

The use of the naphtha dilutent in the aforedescribed manner, i.e., bycracking the naphtha diluent in a secondary or preliminary reactionzone, is desirable from other standpoints in addition to the beneficiallowering of the hydrocarbon feed partial pressure. Up to about 8Opercent of the diluent naphtha can be cracked selectively in thesecondary reaction zone without impairing the reaction in the primarycracking zone. In our novel cracking operation there is an unexpectedand significant upgrading of a portion of the naphtha or gasolinematerial to more desirable gasoline constituents, which unexpectedlyraises the octane rating of that portion of the gasoline product. Wedeem this to be a surprising result as conventional cracking operationsheretofore have produced gasoline from larger non-gasoline moleculesrather than an upgrading of existing gasoline constituents. Asignificant and increased production of C3, C4 and C5 olens results fromthe cracking operation, according to my present invention, which can besubsequently alkylated to provide a further quantity of desirablegasoline constituents. Although a lesser quantity of naphtha diluentdesirably is used, we have successfully employed as much as about 45percent by volume of diluent naphtha, based on the combined initialliquid charge to the primary and secondary reaction zones, and haveachieved unexpectedly advantageous conversion rates, gasoline yield,octane rating, and olefin production. On the other hand, the amount ofcoke deposited upon the fluidized catalyst by cracking of the diluentnaphtha in the preliminary or secondary reaction zone was less than 0.1percent by weight with a diluent naphtha proportion as high as about 45percent.

A zeolite catalyst is a highly suitable catalytic material for use withthis invention. A mixture of natural and synthetic zeolites can beemployed. Also a mixture of crystalline zeolitic organosilicates withnon-zeolitic amorphous Size (microns) Wt. percent -20 0-5 45 20-30 45-7535-55 75 20-40 These particle sizes are usual and have not beenpreselected for this invention. A suitable weight ratio of catalyst toprimary charge is about 4:1 to about 12:1 or 15:1 or even :1, generally;or 6:1 to 10:1, preferably. On the other hand the weight ratio ofcatalyst to naphtha diluent can vary between about 15:1 and about 100:1.The fresh hydrocarbon feed is generally preheated to a temperature ofabout 600 F. to 700 F. but is generally not vaporized during preheat,and the additional heat required to achieve the desired reactortemperature is imparted by the still hot, regenerated catalyst and addeddiluent, issuing from the seconary reaction zone.

The weight ratio of catalyst to hydrocarbon charge is varied to affectvariations in reactor temperature. Furthermore, the higher thetemperature of the regenerated catalyst the less catalyst is required toachieve a given reaction temperature. Therefore, a high regeneratedcatalyst temperature will permit the very low reactor density level setforth below and thereby help to avoid backmixing in the reactor.Generally, catalyst regeneration can occur at an elevated temperature ofabout 1240 F. or 1250 F. or more to reduce the level of carbon on theregenerated catalyst from about 0.6 to 1.5 to about 0.05 to 0.3 percentby weight. At usual catalyst to oil (naphtha and gas oil) ratios thequantity of catalyst is more than ample to achieve the desired catalyticeffect, in both the primary and secondary reaction zones, and thereforeif the temperature of the catalyst is high, the ratio can be safelydecreased without impairing conversion. Since zeolitic catalysts areparticularly sensitive to the quantity of carbon deposited thereon,regeneration advantageously occurrs at elevated temperatures in order tolower the carbon level on the catalyst to the stated range or lower.Moreover, since an important function of the catalyst is to contributeheat to the reactor, for any given desired series of reaction zonetemperatures the higher the temperature of the catalyst charge the lesscatalyst is required, the lower the catalyst charge rate, and the lowerthe density of the material in the reaction zones. As stated, lowreaction zone densities help to avoid backmixing.

The reactor linear velocity, while not being so high that it inducesturbulence and excessive backmixing, must be suciently high thatsubstantially no catalyst accumulation or build-up occurs in eitherreaction zone because such accumulation itself leads to backmixing.Therefore, the catalyst to hydrocarbon weight ratio at any positionthroughout each of the reaction zones desirably is maintained about thesame. Stated another way, catalyst and hydrocarbon at any linearposition along the reaction path in each cracking zone both flowconcurrently at about the same linear velocity, thereby avoidingsignificant slippage of `:atalyst relative to the hydrocarbon component.A build-up of catalyst in either reaction zone leads to a dense bed andbackmixing which in turn increases the residence time in that zone forat least a portion of the charge and induces aftercracking. Avoiding acatalyst build-up in the reaction zones results in a minimal catalystinventory in the reactor, which in turn results in a high spacevelocity. Therefore, a space velocity of over 100 or 120 weight ofprimary hydrocarbon feed per hour per weight of catalyst inventory andabout 200 to 2000 (normally around 1000) weight of diluent naphtha feedper hour per Weight of catalyst inventory, is highly desirable. In theprimary reaction zone the space velocity should not be below 35 and'canbe as high as 500 with reference to the combined gas oil and naphthadiluent. Owing to the low catalyst inventory and low charge ratio ofcatalyst to total hydrocarbon, the density of the material at the inletof the primary reaction zone where the feed is charged can be as low asabout l to less than 5 pounds per cubic foot, although these ranges arenonlimiting. An inlet density in the secondary zone, where the diluentnaphtha and catalyst is charge, below 4 or 4.5 pounds per cubic foot isdesirable since this density range is too low to encompass dense bedsystems, which induce backmixing. Although conversion falls oif with adecrease in inlet density to very low levels, the extent ofaftercracking is a more limiting feature than total conversion of freshfeed, even at an inlet density of less than 4 pounds per cubic foot. Atthe outlet of either reaction zone the density of the corresponding uidstream will be about half the density at the inlet because the crackingoperation in either the naphtha diluent or the gas oil charge producesabout a fourfold increase in mols of gaseous hydrocarbons. The decreasein density through either reaction zone can be a measure of the relatedconversion.

A wide variety of hydrocarbon oil charge stocks can be employed. Asuitable primary charge is a gas oil boiling in the range of 430 F. to1l00 F. As much as 5 to 20 percent of the fresh charge can boil abovethis range. Some residual oil can be charged. A zero to 5 percentrecycle rate can be employed. Generally, the recycle Will comprise atleast 650 F. oil from the product distillation zone which containscatalyst slurry. If there is no catalyst entrainment, recycle can beomitted.

Tests have been conducted, as set forth in detail in the aforementionedcopending application, to illustrate the advantage of the crystallinezeolite aluminosilicate catalyst over an amorphous silica-aluminacatalyst in a iluid catalytic cracking system. The zeolite catalystsystem exhibited a higher conversion rate (85.5 percent compared to 75.5percent) and a higher gasoline yield (61.0 percent to 47.5 percent).However, while the total yield of C3 and C4 hydrocarbons is about thesame for the zeolite and the amorphous catalyst systems, the proportionof C3 and C4 hydrocarbons which are olenic is lower when utilizing azeolite catalyst. As noted previously, this represents a disadvantageousfeature of the zeolite catalyst because C3 and 'C4 oleiins are usefulfor the production of alkylate which can be blended with the gasolineproduced directly by cracking to improve its octane rating. As mentionedpreviously, the disadvantageous lessening of olefin production withzeolite catalyst systems is more than counterbalanced by light olefnicyields from cracking of the naphtha diluent.

A series of tests have been conducted which illustrate the elfect ofhydrocarbon partial pressure (reduced with an inert diluent) uponselectivity to debutanized gasoline and to C3 plus liquid yields. Thecharge stock inspections and other test conditions are detailed in theaforementioned copending application.

The results of the tests are illustrated in FIG. 2 in which debutanizedgasoline yield and total C3 plus liquid yi eld, both recorded as rcentby volume of fresh feed, are plotted against total conversion at variouspartial pressures of hydrocarbon in the system and at various residencetimes. The pressure ranges given on the face of the graphs indicate thepartial pressure in the system of all primary hydrocarbon vapors,cracked and uricracked. For each partial pressure, conversion data isindicated for one or more residence times.

As shown in FIG. 2, at any given conversion level the selectivity togasoline as well as to total C3 plus liquid increases with decreasinghydrocarbon partial pressure. Taking a 60 percent conversion level forpurposes of example, when the hydrocarbon partial pressure is 16-20p.s.i.g., the gasoline yield is 47.5 percent; when the hydrocarbonpartial pressure is l14 p.s.i.g. the gasoline yield increases to almost50 percent; and when the hydrocarbon partial pressure is 2-5 p.s.i.g.the gasoline yield increases still further to about 51.5 percent.Advantageously, a greater improvement in gasoline selectivity occurredin reducing hydrocarbon partial pressure from 16-20 p.s.i.g. to 10-14p.s.i.g. than occurred in reducing the hydrocarbon partial pressure from10-14 p.s.i.g. to the very low partial pressure of 2-5 p.s.i.g.

Example I To demonstrate the efficacy of our invention in upgradingunconverted naphtha, in increasing lgasoline yield and increasingolefinic production, We have run tests utilizing a virgin, parairiicnaphtha as diluent material and a full range gas oil as primaryhydrocarbon charge. The naphtha was charged in suiiicient quantity toevaluate yields and was for test purposes 44.4 volume percent of thetotal charge. The naphtha was preliminarily cracked at 1200" F. for 2seconds in the lower portion of the transfer line of the apparatusdescribed below in connection with FIG. 3, which corresponds to theaforementioned preliminary or secondary reaction zone. The gas oil wascracked at 1000 F. for 0.5 in the riser portion of the apparatus, i.e.,the primary reaction zone. 'I'he cracked naphtha was, of course, addedto the gas oil at the entrance of the primary reaction zone to reducethe partial pressure of the gas oil. The aforementioned pilot plant runwas compared with a similar run using gas oil only, and the completeresults of the two runs are set forth in the following table.

TABLE I Combined Gas oil Run Number feeds I only II Charge stock:

Naphtha charge rate, gmJhr 360 636 Gas oil charge rate, gm./hr 450Catalyst; cracking conditions:

Transfer line temperature, F.:

Top 1,000 1, 000 Bottom 1,200 1, 000 Contact time, seconds:

Gas oil charge 0. 5 2. 5 Naphtha charge. 2.5 Cat/oil ratio, wtJwt.:

0n fresh feed. 8. 2 8. 7 On total feed--. 8. 2 8. 7 Total charge rate,g./hr 810 636 Dispersion steam, lla/1,000 lb. catalyst 6. 0 3.6Stripping steam, 1b./1,000 lb. catalyst 6. 1 7.3 Stripping N2 as steam,lb./1,000 lb. catalyst. 5. 1 6.1 Catalyst circulation rates, gin/hr.:

Calculated catalyst circulation 5, 584 4, 587 Measured catalyst and cokecirculation.- 6, 648 5, 568 Coke-free measured catalyst circulation- 6,615 5, 530 Carbon on spent catalyst, Weight percent percent 0n catalyst0. 501 0. 689 Carbon on regeneration catalyst, weight percent oncatalyst 0. 183 0. 219 Operating conditions, regeneration:

Average temperature, F.. 1, 149 1, 149 Average pressure, p.s.i.g .6 29.6 Miscellaneous data:

Uncorrected weight balance 99. 1 97. 1 Stabilized gasoline, weightpercent of fresh feed 56.0 53. 0 H drocarbon as, W

fied E 27. 1 20.4 Total stabilizer gas, s. 201. 45 131.19 l()i gasgrafit .fls 1. 423 1. 370

ri ac ua 1 a 4cent g. 3 84 4.32 C in a t al RVP asoline,

4centi?.1 2.79 2.81 10 RVP gasoline, weight percen 61.2 56.0 Knockrating [adjusted to 10 Reid Va sure (RVP)]:

Motor octane numbers:

ar 79. 7 83. 9 +3 cc. TEL 86.0 87. 5 Research octane numbe ar 87. 6 95.6 |3 cc. TEL 94.6 99.7 Conversion, volume percent of fresh feed 0- 88 l80 1 gas oil) TABLE I-Contlnued Combined Gas oil Run Number feeds I onlyII Yields, volume percent of fresh feed (corrected to 100 weight percentbalance):

Debutanized gasoline distribution Isopentane n-Pentane PentenesPentenes- Hexanes and heavier..

Butane-Butene Product inspections:

Debutani'zed gasoline, gravity: API Hydrocarbon type analysis, volumeper- Debutanized distribution 74. 8 80. 2 Depropanized distribution 93.9 100. 6 Yield data; percent on fresh feed:

Debutanized gasoline distribution 58. 45 53. 19 Isopentane 5. 05 5. 10n-Pentanc.- 3. 22 0. 49 Pentenes 3. 42 2. 79 Hexanes and heavier 46. 7644. 82

Butane-butene 12. 31 10. 50

3. 87 4. 36 1. 50 1. 33 Butenes 6. 94 4. 81

Gas (C3 and lighter) 12. 38 9. 72 Methane 1. 76 1. 43 Ethane 0. 71 0. 86Ethylene 1. 08 0. 85 Propane.- 2. 88 l. 48 Propyleiie- 5. 88 5. 04

Light catalyst gas oi 3. 67 14. 62

Heavy catalyst gas oil 0. 00 0. 00

Decanted oil 10. 30 7. 42

Saturatie Reid vapor pressure, Micro octane number:

:non wel?? PES5. cedido @alom cammeo :eeuwse 93:75.53 @050cm @moviDistillation, D 86:

Over point, F 116 955 End point, F-.. 379 412 10% at, 145 133 50 192 226304 368 Recovery, percent 99. 0 97. 3 Residue, percent-- 0. 6 0. 8 p 0.4 1. 9

Gravity, PI 20. 2 13.0 Sulfur, weight percent-. 0. 73 0. 92 Viscosity,SUS at F 33. 8 38.5 Carbon residue, D 524 10% Btms., weight percent 0.51 Pour point, F -35 10 Distillation, D 86:

Over point, F End point, F 720 10% at, 486 50% at, F. 536 90% at, F 649Decauted oil:

Gravity, API 14. 8 0. 5 Sulfur, weight percent 0. 97 1. 46 Carbonresidue, Rams. bottom, weight percent 0. 55 5. 50 Distillatiou, vacuumcorrected to 760 nim. g

Ihe results of these tests can be summarized as follows:

(l) The conversion of the naphtha diluent to C3 and C4 gases wasestimated at 32 volume percent.

(2) Selectivity to C3 and C4 oleiinic production is estimated to be 62.5percent based on the naphtha charge.

(3) Improvement in quality or upgrading of the unconverted naphtha wasobtained.

(4) Improvement in the cracking of the gas oil charge was obtained owingto the reduced partial pressure in the primary reaction zone or riser.

The charge to the transfer line or secondary reaction zone of theuidized catalyst cracking plant consisted for the purpose of the testrun of 44.4 volume percent virgin Kuwait naphtha ,(105-290 F.) and 55.6volume percent South Louisiana full-range gas oil.

Inspections of the naphtha charge stock are shown in the followingtable:

TABLE II.-NAPHTHA CHARGE STOCK INSPECTIONS Hydrocarbon analysis, ASTM1956, vol percent:

Paratins Cycloparans Bigycloparafns Alkylbenwnes Benzene Toluene. CaAlkylbemenm Total Distillation, D 86 F:

The inspections of the gas oil charge stock are set forth 1n thefollowing table:

TABLE IIL-GAS OIL CHARGE STOCK INSPECTIONS Rnn Number: I II Charge stock(Southern Louisiana gas oil):

Same as Run No. 1. Characterization factor-.. 11.81 Do. Gravity, API 23.5 Do. Sulfur, GRM 1156, wt. percent-. 0. 58 Do. Viscosity, SUS at, F.:

130-... 160 91.0 Do. 210--.- 210 48.9 Do. Carbon residue, Rams., ASTM D524, Wt. percent 0. 19 D0. Aniline point, GRM 139, F.- 186 D0. Brominenumber, D 1159-.-- 3.31 Do. Pour point, D 97, F 76 Do. Refractive index,GRM 2413, nd at 20 C- 1. 5095 Do. Nitrogen, ppm.:

Total, GRM 1121 640 D0: Basic, GRM 1152 237 Do. Metals, p.p.m.:

Vanadium, GRM 803 0. 1 Do. Nickel, GRM 803 0.1 DO. Distillation, vac.(corr. to 760 mm. Hg):

Over point, F Do. Over at, F:

10% 688 D0. 745 Do. 805 D0. 873 Do. 944 Do. Catalyst Kellogg aetivity53. 5 Do. Carbon factor (C) 0. 53 D0. Carbon factor (H)- 0. 62 Do.Hydrogen factor.-... 1. 59 Do.

The naphtha stock was charged to the bottom injector of the transferline, as shown 1n FIG. 3, and cracking occurred at 1200 F., utilizing inthis case a 25:7 catalystoil ratio, and a two-second residence time. Thegas oil, on the other hand, was charged to the bottom of the riser line(as shown in FIG. 3), and cracking occurred in the gas oil and admixedeluent from the preliminary reaction zone at 1000" F., with an 8:2catalyst-oil ratio, and a 0.5 second residence time. The 0.5 secondresidence time wa's less than optimum and resulted primarily from therather larger proportion of added diluent. In the control run, in whichthe charge consisted of gas oil only, cracking occurred at 1000 F., withan 8:7 catalyst-oil ratio and a residence time of 2.5 seconds.

A comparison of the results of the combined naphtha gas oil run and thegas-oil only run are summarized in the following table:

TABLE IV.-NAPHTHAGAS OIL VERSUS GAS OIL CRACKING COMPARISON Naphtha- Gasoil gas oil run run Cracking conditions Charge stock (l) (2) Percent oftotal charge.- 44. 4 100.0 Temperature, F 1, 200 1, 000 Cat./oil ratio,wt./hr./wt 25. 7 8. 7 Contact time, seconds.. 2.0 2. 5 Charge stock (2)Percent oi total charge 55. 6 Temperature, F 1, 000 Cat./oil ratio,wtJhL/wt.- 8.2 Contact time, second 0. 5 Conversion, vol. percent:

13. 8 11.5 9.1 8.8 16. 9 16. 4 9. 2 7. 2 65. 9 64. 2 108. 4 112.0 C2 andlighter, wt. p 3. 6 3. 2 Coke, wt. percent 2. 9 4. 5

1 Kuwait light naphtha.

l Southern Louisiana gas oil.

a Conversion to 430 F. and lighter. 4 Conversion to C4 and lighter.

5 Conversion to C; and lighter.

From the preceding table a signiiicant increase in C3 and C4 olefms canbe noted for the naphtha-gas oil run. It is noted also that the crackingof the ga's oil charge in the experimental run was enhanced, in view ofan estimated conversion of 78.6 percent at a 0.5 second residence timein the naphtha-gas oil run, which compares very favorably with an 80.1percent conversion at a considerably longer and more nearly optimumresidence time of 2.5 seconds in the control run.

In the following table the gasoline product quality of the experimentalnaphtha-gas oil run is compared to a typical gasoline quality obtainedfrom the control run:

As noted from the preceding table, gasoline quality in the experimentalrun decreased somewhat. This was not unexpected, as a virgin naphthafraction (50-75 octane) was employed in the experimental naphtha-gas oilrun and moreover comprised nearly half of the total feed stock.

Most significantly, however, the gasoline octane ratings decreased onlyin the range of 2-8.5 octane numbers. It must be remembered that naphthacharge was predominately saturates (96 percent). It follows, then, thatthe gasoline composition resulting from the experimental run should havebeen drastically changed if the quality of the unconverted naphtha hadnot been improved. It was to be expected, for example in the case of themotor +3 and research +3 octane ratings, that the results of theexperimental run should have about midway between the naphtha aloneratings (75 octane) and the cracked gas oil alone ratings, or about 81.0and A87.0 respectively, assuming a maximum naphtha octane rating of 75.Instead, the actual ratings as seen from Table V were 85.1 and 93.3respectively, which denote a most significant and marked upgrading ofthe unconverted portion of the naphtha charge. This considerableincrease of at least 4 and 6 octane points in the motor +3 and research-l-3 ratings, would alone save respectively a minimum of 2 and 3 gms. oflead additives per gallon of gasoline. Further evidence of thebeneficial results flowing from the use of naphtha as a diluent can begleaned from Table V based upon a percent increase in oleinic content, adecrease of 12 percent in aromatic content and in increased saturatecontent of only 7 percent in the gasoline product. These comparisonsfurnish a further indication of significant improvement or upgrading ofthe naphtha diluent.

Utilizing the yield data shown for the gas oil run (Tables I and IV) thefollowing table presents the yield and quality estimates for thecracking of a virgin naphtha charge alone:

TABLE VI.CRACKING OF LIGHT KUWAIT NAPHTHA, ESTIMATED YIELDS Crackingconditions:

Transfer line temperature, F Cat-to-oil ratio Contact time, secondsProduct distribution, vol. percent:

Total C1 4= Total C1 C Conversion of the naphtha charge to C3 and C4light ends is calculated to be 32.0 volume percent. At this conversionrate, a 62.5 percent selectivity to C3 and C4 olens was achieved. It isfurther estimated that a 50.7 percent conversion of the naphtha chargeto C3, C4, and C5 gases was obtained. The overall quality of the naphthacharge is therefore shown to be substantially improved even at thisconversion level.

A suitable reactor-regenerator system for performing our invention isdescribed with reference to FIG. 3. The cracking of the gas oil in thecombined charge occurs in the primary reaction zone which includes anelongated reactor tube 10, usually referred to as a riser. In thisexample, the riser has a length to diameter ratio of above 20 or 25. Afull range hydrocarbon oil feed to be cracked is passed throughpreheater 2 to heat it to about 600 F. and is then charged into thebottom of the riser 10 through an inlet line 14. Steam or other inertdiluent, if desired, can be introduced into the oil inlet line 14through inlet 18.

Similarly, steam or other inert diluent can be introduced independentlyand directly to the primary reaction zone, i.e. to the bottom of theriser 10 through line 22, where desired for minor adjustments in partialpressure, residence time, catalyst uidization, etc. Depending upon theamount of naphtha diluent added as described below, such inert diluent,for example, can aid in carrying upwardly into the riser 10 theregenerated catalyst stream which ows to the bottom of the riser 10through transfer line 26.

The preponderate proportion or all of the diluent for lowering thepartial pressure of the gas oil is added, how ever, to the transfer-line 26 through inlet line 27 at a predetermined distance from thejunction between the transfer line 26 and the riser 10, which definesthe secondary or preliminary reaction zone. The amount of naphthadiluent added through the inlet 27 can vary from about 5 volume percentbased on the gas oil charge to about 45 percent or more. The boilingrange of the naphtha diluent can be selected as described previously,and the catalyst to oil ratio both in the transfer line 26 and in theriser 10 can be adjusted as required by means of valves 40, 41 and 42.It will be seen from FIG. 3 that the naphtha diluent is addedsufliciently upstream of the riser 10, in this example, to achieve aprescribed contact time or residence time of the naphtha and catalyststreams within the preliminary reaction zone. In this case, theresidence time is about 2 seconds although considerable variation ispossible depending upon a specic application of the invention. Up toabout percent of the naphtha can be converted in the secondary reactionzone, i.e. in the transfer line 26, and as the naphtha is morerefractory than the gas oil, substantially all of the naphtha conversiontakes place in the transfer line. It will be understood, of course, thatboth the naphtha diluent and the gas oil can be added directly to theriser 10, for example through alternative inlets 28, 29 respectively.The distance between the entry points of the inlets 28, 29 woulddetermine the preliminary or secondary reaction zone.

The gas oil to be cracked in the riser 10 desirably has a boiling rangeof about 430 to 1100 F. The catalyst employed is a liuidized zeoliticaluminosilicate and is introduced into the riser 10 adjacent the bottomthereof where the riser is adjoined with the descending transfer line26. Depending upon the particular boiling range of a specific gas oilcharge, the riser temperature is maintained within the range of about900-1100 F. and preferably within the range of 950'-1000. The risertemperature is controlled by measuring the temperature of the productfrom the riser and then by adjusting the opening of valve 40 by means oftemperature controller 43 to regulate the inflow of hot regeneratedcatalyst through transfer line 26.

'Ihe temperature of theregenerated catalyst as it flows from theregenerator 64 into the transfer line 26 is considerably above thecontrol temperature in the riser 10 or primary reaction zone so that theincoming catalyst contributes heat to the cracking reactions of thenaphtha diluent in the lower portion of the transfer line 26 and as wellas the gas oil in the riser 10. The riser pressure desirably is in therange of about 1035 p.s.i.g. Between about 0 and 5 percent of the oilcharge tothe riser 10 can be recycled (not shown). The residence time ofthe gas oil, converted naphtha and catalyst in the riser 10 is verysmall and ranges from about 0.5 to 5 seconds. The residence time in theprimary reaction zone or riser 10 usually is shorter than in thesecondary reaction zone or lower portion of the transfer line 26. Thevelocity of the catalytic stream through the apparatus is about 35-55ft. per second in order to minimize or prevent altogether any slippagebetween the hydrocarbon and catalyst, particularly in the riser 10.Therefore, no bed of catalyst is permitted to build up throughout theapparatus, and in furtherance of this purpose, the density within theriser 10 is a very low maximum of about four pounds per cubic foot atthe bottom of the riser and decreases to about two pounds per cubic footat the top of the riser. Since no dense bed of catalyst is permitted tobuild up within the transfer line 26 and riser 10, the space velocitythrough the apparatus is unusually high and will have a range between100 and 120 and 600 weight of hydrocarbon per hour per instantaneousweight of catalyst in the reactor. No significant catalyst build upwithin the reactor is permitted to occur and the instantaneous catalystinventory within the lower portion of the transfer line 26 and in theriser 10 is due to a flowing catalyst to oil weight ratio in the rangeof about 4:1-l5:l, the weight ratio corresponding to the feed ratio.

The hydrocarbon and catalyst exiting from the top of the riser is passedinto a disengaging vessel 44. The top of the riser is capped at 46 sothat discharge occurs through lateral slots 50 for proper dispersion. Aninstantaneous separation between hydrocarbon and catalyst occurs in thedisengaging vessel, which terminates the cracking reaction. Thehydrocarbon which separates from the catalyst is primarily gasolinetogether with some heavier components and some lighter gaseouscomponents. The hydrocarbon efuent passes through a cyclone system S4 toseparate catalyst fines contained therein and is discharged to afractionator through line 56. The catalyst separated from hydrocarbon inthe disengager 44 immediately drops below the outlets of the riser sothat there is no catalyst level in the disengager but only in a lowerstripper section 58. Steam is introduced into catalyst stripper section58 through sparger 60 to remove any entrained hydrocarbon in thecatalyst.

Catalyst leaving stripper 58 passes through transfer line 62 to aregenerator 64. This portion of the catalyst contains carbon depositswhich tend to lower its cracking ecacy and as much carbon as possiblemust be burned from the surface of the catalyst. As noted previously,virtually all of the carbon deposit is derived from the gas oil portionof the total hydrocarbon charge.

Burning is accomplished by introduction to the regenerator through line66 of approximately the stoichiometrically required amount of air forcombustion of the carbon deposits. The catalyst from the stripper entersthe bottom section of the regenerator in a radial and downward directionthrough transfer line 62. Flue gas leaving the dense catalyst bed inregenerator 64 flows through cyclones 72 wherein catalyst fines areseparated from flue gas permitting the flue gas to leave the regeneratorthrough line 74 and pass through a turbine 76 before leaving for a wasteheat boiler wherein any carbon monoxide contained in the flue gas isburned to carbon dioxide to accomplish heat recovery. Turbine 76compresses atmospheric air in air compressor 78 and this air is chargedto the bottom of the regenerator through line 66.

The temperature throughout the dense catalyst bed in the regenerator isin the neighborhood of 125()D F., and preferably is maintained about 250F. above the control temperature in riser 10. The temperature of theflue gas leaving the top of the catalyst bed in the regenerator can risedue to afterburning of carbon monoxide to carbon dioxide. Approximatelya stoichiometric amount of oxygen is charged to the regenerator and thereason for this is to minimize afterburning of carbon monoxide to carbondioxide above the catalyst bed to avoid injury to the equipment since atthe temperature of the regenerator flue gas some afterburning doesoccur. In order to prevent excessively high temperatures in theregenerator ue gas due to afterburning, the temperature of theregenerator flue gas is controlled by measuring the temperature of theflue gas entering the cyclones and then venting some of the pressurizedair otherwise destined to be charged to the bottom of the regeneratorthrough vent 80 in response to this measurement. The regenerator reducesthe carbon content of the catalyst from 110,5 weight percent to 0.2weight percent or less. If required, steam is 22 available through line82 for cooling the regenerator. Makeup catalyst is added tothe bottom ofthe regenerator through line 84. Hopper S6 is disposed at the bottom ofthe regenerator for receiving regenerated catalyst to be passed to thebottom of the reactor riser through transfer line 26.

With cracking of the naphtha diluent and with a significant proportionof the gasoline product being derived from the naphtha diluent, it isevident that the partial pressure of the gas oil feed has beensignificantly and advantageously lowered. The advantages of such furtherlowering of the partial pressure of the gas oil feed is evident fromFIG. 2 of the drawings, as noted previously.

From the foregoing it will be apparent that novel and efficient forms ofa Fluid Catalytic Cracking Process have been described herein. While wehave shown and described certain presently preferred embodiments of theinvention and have illustrated presently preferred methods of practicingthe same, it is to be distinctly understood that the invention is notlimited thereto but may be otherwise variously embodied and practicedwithin the spirit and scope of the invention.

We claim:

1. .A process for a lowered partial pressure cracking of a principalhydrocarbon charge in the presence of a stream of fluidized zeolitecracking catalyst to obtain a higher selectivity for and/or octanerating of gasoline constituents and increased production of C3, C4 andC5 olefins, said process comprising the steps of maintaining apredetermined range of temperatures within said zeolite catalyst stream,adding to said catalyst stream a naphtha diluent boiling between aboutF. and about 290 P. and having about 7096% saturates, adding saidprincipal charge to said catalyst stream, maintaining a lower partial rpressure of said charge in said stream by maintaining a given ratio ofsaid diluent to said charge, adding said naphtha diluent to saidcatalyst stream at a point having a higher temperature than that atwhich said charge is added so that a significant proportion of each ofsaid naphtha and said charge can be cracked by said catalyst, locatingsaid naphtha addition sufficiently upstream of the principal chargeaddition so that substantially all of the cracking and/or conversion ofsaid naphtha takes place prior to said charge addition, thereafteradmixing said diluent including the cracked products thereof with saidcharge to increase the effective volume of said diluent and to lowerfurther the partial pressure of said charge, limiting the residence timeof said principal charge in said zeolite catalyst stream to a maximum ofabout ve seconds to avoid masking of said higher selectivity byaftercracking of said charge and of said naphtha, and limiting the totalresidence time of said naphtha in said zeolite catalyst stream to amaximum of about 20 seconds to avoid aftercracking and attendantpolymerization of said constituents and said cracked products and tominimize coking of said catalyst and the production of C2 and lightergases.

2. The process according to claim 1 including the modified step ofadding said naphtha diluent to said catalyst stream at a point having atemperature of the order of about 250 F. higher than that at which saidprincipal charge is added.

3. The process according to claim 1 including the modified step ofadding'said naphtha diluent to said catalyst stream at a locationupstream of said principal charge addition to define a preliminarycracking zone for said diluent and a primary cracking zone communicatingtherewith for receiving directly therefrom said catalyst, diluent andcracked diluent products, said principal charge being added directly tosaid primary zone.

4. The process according to claim 3 including the additional steps ofmaintaining a higher temperature and a given residence time in saidpreliminary cracking zone,

23 and maintaining a lower temperature and a different residence time insaid primary cracking zone.

5. The process according to claim 3 including the additional step ofadding a minor proportion of said naphtha diluent directly to theentrance of said primary zone.

6. The process according to claim 1 including the modified step ofadding said naphtha diluent in the range of about 5 percent to about 45percent by volume of the total hydrocarbon feed.

7. The process according to claim 1 including the additional step ofestablishing a ratio of said naphtha diluent and the conversion productsthereof to said principal charge and establishing residence time of saidprincipal charge such that a greater percentage yield of gasoline basedon total hydrocarbon feed is recovered from said process in the presenceof said naphtha diluent and its conversion products than could berecovered from said process in the absence of said naphtha diluent.

8. The process according to claim 1 including the additional step ofestablishing the ratio of said naphtha diluent and its conversionproducts to said principal charge such that a greater percentage yieldof lighter oletins 'based on total hydrocarbon feed is recovered fromsaid process in the presence of said -naphtha diluent than could berecovered from said process in the absence of said diluent.

9. The process according to claim 1 including the additional step ofestablishing the ratio of said naphtha diluent to said principal chargeand a predetermined temperature and residence time of said diluent insaid catalyst stream prior to engagement with said principal charge suchthat the octane rating of the unconverted proportion of said naphthadiluent is upgraded.

10. The process according to claim 1 including the addition step ofadding a minor quantity of said diluent directly to said principalcharge prior to its introduction into said catalyst stream.

11. The process according to claim 1 including the modified step ofestablishing said naphtha residence time within the range of about 2 to20 seconds.

12. The process according to claim `11 including the further modifiedstep of limiting said residence time to about 2-10 seconds.

13. The process according to claim 1 including the additional step ofmaintaining a catalyst to naphtha ratio during said naphtha residencetime of between about 15 :1 and about 100:1.

14. The process according to claim 1 including the additional step ofmaintaining the density of the combined catalyst and naphtha streambelow about 4.5 pounds per cubic feet.

15. The process according to claim 1 including the additional step ofmaintaining a space velocity relative to said naphtha residence timefrom about 200 to about 2000 weight of diluent naphtha feed per hour.

16. The process according to cla-im 15 including the modified step ofestablishing saiid space velocity in the 24 neighborhood of 1000 weightof diluent naphtha feed per hour.

17. The process according to claim 1 including the modified step ofadding said naphtha diluent having at least by volume of saturates.

18. The process according to claim 1 including the modied step of addingsaid naphtha diluent having at most 10% by volume of olelins.

19. The process according to claim 1 including the modified steps ofcracking said charge in a riser zone, and disengaging said charge andthe cracked products thereof from said catalyst stream at the exit ofsaid riser zone to avoid bed-cracking in a subsequent separation zone.

20. The method according to claim 1 including the modified step oflimiting the residence time of said charge to a maximum of about 5seconds.

21. The process according to claim 1 wherein said naphtha diluentincludes that fraction thereof having the lowest octane rating.

22. The process according to claim 1 wherein the residence time of saidpnincipal charge is Within about 0.5 to about 4 seconds, and theresidence time of said diluent is within about 2 seconds to about 10seconds.

23. The process according to claim 1 wherein the velocity of saidprimary charge in said zeolite catalyst stream is within about 25 toabout 75 feet per second.

24. The process according to claim 1 wherein said principal charge isadded adjacent the bottom of a riser cracker through which said catalyststream is circulated, and said diluent is added to a transfer conduitcoupling the bottom of said riser to a catalyst regenerator.

25. The combination according to claim 1 wherein said process is carriedout in cracking apparatus including a riser tube and a transfer lineconnected thereto, said charge being added to said riser tube with atleast a portion thereof added adjacent the junction thereof with saidtransfer line, and said naphtha is added to said transfer line at apoint upstream of said junction.

References Cited UNITED STATES PATENTS 2,890,164 6/ 1959 Woertz 208-742,893,943 7/ 1959 Vignovich 20S-78 2,908,630 10/1959 Friedman 208-153 X3,042,196 7/ 1962 Payton et al. 208-113 3,158,562 11/1964 Peet 20S-153 X3,186,805 6/ 1965 Gomory 20S-153 X 2,921,014 1/ 1960 Marshall 20S-743,424,672 1/ 1969 Mitchell 208-164 3,448,037 `6/ 1969 Bunn, Jr., et al20S-74 3,617,497 11/1'971' Bryson et al. 20S-74 DELBERT' E. GANTZ,Primary Examiner U.S. Cl. X.R.

